The present invention relates to the field of catalytic oxidation of hydrocarbons. More particularly, the present invention relates to the catalytic partial oxidation of paraffinic hydrocarbons, such as ethane, propane, and naphtha, to produce olefins, such as ethylene and propylene.
Olefins find widespread utility in industrial organic chemistry. Ethylene is needed for the preparation of important polymers, such as polyethylene, vinyl plastics, and ethylene-propylene rubbers, and important basic chemicals, such as ethylene oxide, styrene, acetaldehyde, ethyl acetate, and dichloroethane. Propylene is needed for the preparation of polypropylene plastics, ethylene-propylene rubbers, and important basic chemicals, such as propylene oxide, cumene, and acrolein. Isobutylene is needed for the preparation of methyl tertiary butyl ether. Long chain mono-olefins find utility in the manufacture of linear alkylated benzene sulfonates, which are used in the detergent industry.
Low molecular weight olefins, such as ethylene, propylene, and butylene, are produced almost exclusively by thermal cracking (pyrolysis/steam cracking) of alkanes at elevated temperatures. An ethylene plant, for example, typically achieves an ethylene selectivity of about 85 percent calculated on a carbon atom basis at an ethane conversion of about 60 mole percent. Undesired coproducts are recycled on the shell side of the cracking furnace to be burned, so as to produce the heat necessary for the process. Disadvantageously, thermal cracking processes for olefin production are highly endothermic. Accordingly, these processes require the construction and maintenance of large, capital intensive, and complex cracking furnaces. The heat required to operate these furnaces at a process temperature of about 900xc2x0 C. is frequently obtained from the combustion of methane which disadvantageously produces undesirable quantities of carbon dioxide. As a further disadvantage, the crackers must be shut down periodically to remove coke deposits on the inside of the cracking coils.
Catalytic processes are known wherein paraffinic hydrocarbons are oxidatively cracked to form mono-olefins. In these processes a paraffinic hydrocarbon is contacted with oxygen in the presence of a catalyst consisting of a platinum group metal or mixture thereof deposited on a ceramic monolith support. Optionally, hydrogen may be a component of the feed. The process is conducted under autothermal reaction conditions wherein the feed is partially combusted, and the heat produced during combustion drives the endothermic cracking process. Consequently, under these autothermal process conditions there is no external heat source required; however, the catalyst is required to support combustion above the normal fuel-rich limit of flammability. Representative references disclosing this type of process include the following U.S. Pat. Nos. 4,940,826; 5,105,052; 5,382,741; and 5,625,111. Disadvantageously, substantial amounts of deep oxidation products, such as carbon monoxide and carbon dioxide, are produced, and the selectivity to olefins remains too low when compared with thermal cracking.
M. Huff and L. D. Schmidt disclose in the Journal of Physical Chemistry, 97, 1993, 11,815, the production of ethylene from ethane in the presence of air or oxygen under autothermal conditions over alumina foam monoliths coated with platinum, rhodium, or palladium. A similar article by M. Huff and L. D. Schmidt in the Journal of Catalysis, 149, 1994, 127-141, discloses the autothermal production of olefins from propane and butane by oxidative dehydrogenation and cracking in air or oxygen over platinum and rhodium coated alumina foam monoliths. The olefin selectivity achieved in these processes is not comparable to that achieved by steam cracking and therefore could be improved.
U.S. Pat. No. 5,639,929 teaches an autothermal process for the oxidative dehydrogenation of C2-C6 alkanes with an oxygen-containing gas in a fluidized catalyst bed of platinum, rhodium, nickel, or platinum-gold supported on alpha alumina or zirconia. Ethane produces ethylene, while higher alkanes produce ethylene, propylene, and isobutylene. Again, the olefin selectivity could be improved.
C. Yokoyama, S. S. Bharadwaj and L. D. Schmidt disclose in Catalysis Letters, 38, 1996, 181-188, the oxidative dehydrogenation of ethane to ethylene under autothermal reaction conditions in the presence of a bimetallic catalyst comprising platinum and a second metal selected from tin, copper, silver, magnesium, cerium, lanthanum, nickel, cobalt, and gold supported on a ceramic foam monolith. This reference is silent with respect to co-feeding hydrogen in the feedstream. While the use of a catalyst containing platinum and tin and/or copper is better than a catalyst containing a platinum group metal alone, the olefin selectivity should be improved if the process is to be commercialized.
In view of the above, it would be desirable to discover a catalytic process wherein a paraffinic hydrocarbon is converted to an olefin in a conversion and selectivity comparable to commercial thermal cracking processes. It would be desirable if the catalytic process were to produce only small quantities of deep oxidation products, such as, carbon monoxide and carbon dioxide. It would also be desirable if the process were to achieve low levels of catalyst coking. It would be even more desirable if the process could be easily engineered without the necessity for a large, capital intensive, and complex cracking furnace. Finally, it would be most desirable if the catalyst for the process exhibited good stability.
This invention is a process for the partial oxidation of paraffinic hydrocarbons to form olefins. The process comprises contacting a paraffinic hydrocarbon or mixture thereof with oxygen in the presence of hydrogen and a catalyst. The contacting is conducted under autothermal process conditions sufficient to form the olefin. The catalyst employed in the process of this invention comprises a Group 8B metal and at least one promoter.
The process of this invention efficiently produces olefins, particularly mono-olefins, from paraffinic hydrocarbons, oxygen, and hydrogen. Advantageously, the process of this invention achieves a higher paraffin conversion and a higher olefin selectivity as compared with prior art catalytic, autothermal processes. More advantageously, the process of this invention produces fewer undesirable deep oxidation products, such as carbon monoxide and carbon dioxide, as compared with prior art catalytic, autothermal processes. Even more advantageously, in preferred embodiments, the process of this invention achieves a paraffin conversion and olefin selectivity which are comparable to commercial thermal cracking processes. As a further advantage, the process produces little, if any, coke, thereby substantially prolonging catalyst lifetime and eliminating the necessity to shut down the reactor to remove coke deposits.
Most advantageously, the process of this invention allows the operator to employ a simple engineering design and control strategy, which eliminates the requirement for a large, expensive, and complex furnace like that used in thermal cracking processes. In one preferred embodiment, the reactor simply comprises an exterior housing which contains a monolithic support onto which the catalytic components are deposited. Since the residence time of the reactants in the process of this invention is on the order of milliseconds, the reaction zone operates at high volumetric throughput. Accordingly, the reaction zone measures from about one-fiftieth to about one-hundredth the size of a commercially available steam cracker of comparable capacity. The reduced size of the reactor lowers costs and simplifies maintenance procedures. Finally, since the process of this invention is exothermic, the heat produced can be harvested via integrated heat exchangers to generate electricity or steam credits for other processes.
As noted hereinbefore, thermal energy is needed to maintain autothermal process conditions. Without preheating the feedstream, the required thermal energy is totally supplied by the reaction of the feedstream with oxygen, namely, alkane oxidative dehydrogenation to form olefins and water, hydrogen oxidation to form water, and carbon combustion to form carbon monoxide and carbon dioxide. These processes can supply the heat necessary for any endothermic dehydrogenation which takes place to form ethylene and hydrogen. The prior art has recognized that a portion of the required thermal energy can be obtained by preheating the feedstream. The preheat can be conveniently supplied by condensing high pressure saturated steam, or alternatively, by combusting process off-gas or another fuel source. Surprisingly, it has now been discovered that a high preheat temperature can be used without loss in olefin selectivity, and further, that a high preheat temperature provides advantages unrecognized heretofore. Accordingly, in another aspect of this invention, the paraffinic hydrocarbon and oxygen, which together comprise the reactant feedstream, are preheated at a temperature greater than about 200xc2x0 C., but below the onset of reaction of the feedstream components.
When the high preheat temperatures of this invention are employed, advantageously less oxygen is required in the feedstream. Since the cost of pure oxygen can be a significant cost component of the feedstream, the decrease in oxygen employed translates directly into economic savings. Moreover, since oxygen reacts with hydrogen in the feedstream, the decrease in oxygen employed leads to a decrease in hydrogen consumed and in the waste water produced. As a consequence, more hydrogen is found in the product stream.
An increased yield of hydrogen in the product stream further improves the economics of the autothermal oxidation process of this invention. Since hydrogen is required for the process, hydrogen in the product should be recycled and any deficit must be replaced by importing hydrogen from an external source. Alternatively, hydrogen can be made from off-gas streams, for example, a water-shift reaction which converts carbon monoxide and water to hydrogen and carbon dioxide. As a consequence of using the high preheat temperature of this invention, the product stream is enriched in hydrogen. Under optimal preheat conditions, the recycled hydrogen substantially eliminates the need to import hydrogen or to derive make-up hydrogen from other sources.
In a third aspect, the autothermal oxidation process of this invention is beneficially conducted in a unique fluidized bed reactor, characterized in that the reactor bed possesses an aspect ratio of less than about 1:1, as measured during operation. For the purposes of this invention, the aspect ratio is defined as the ratio of the height (or depth) of the reactor bed to its cross-sectional dimension (diameter or width). For use in this fluidized bed, the catalyst comprises a support in the form of pellets or spheres onto which the catalytic components are deposited.
When operation of the process in the aforementioned unique fluidized bed reactor is compared with operation in a fixed bed reactor, several advantages become apparent. For example, ethylene selectivity improves with use of the fluidized bed, while selectivities to methane and deep oxidation products, such as carbon monoxide and carbon dioxide, decrease. Significantly, the selectivity advantages are achieved at ethane conversions which are comparable to or better than those obtained in a fixed bed reactor.
In a fourth aspect, this invention is a catalyst composition comprising a Group 8B metal and at least one promoter supported on a catalyst support which has been pretreated with at least one support modifier.
The aforementioned composition is beneficially employed as a catalyst in the autothermal partial oxidation of a paraffinic hydrocarbon to an olefin. The catalyst composition beneficially produces an olefin or mixture of olefins at conversions and selectivities which are comparable to those of industrial thermal cracking processes. Accordingly, the catalyst composition of this invention produces low amounts of carbon monoxide and carbon dioxide. Finally, the catalyst composition of this invention advantageously exhibits good catalyst stability.
The process of this invention involves the partial oxidation of a paraffinic hydrocarbon to form an olefin. The words xe2x80x9cpartial oxidationxe2x80x9d imply that the paraffin is not substantially oxidized to deep oxidation products, specifically, carbon monoxide and carbon dioxide. Rather, the partial oxidation comprises one or both of oxidative dehydrogenation and cracking to form primarily olefins. It is not known or suggested to what extent or degree either process, oxidative dehydrogenation or cracking, predominates or occurs to the exclusion of the other.
The partial oxidation process of this invention comprises contacting a paraffinic hydrocarbon with oxygen in the presence of a multi-metallic catalyst and in the presence of a hydrogen co-feed. The contacting is conducted under autothermal process conditions sufficient to form the olefin. The catalyst which is employed in the process of this invention comprises a Group 8B metal and at least one promoter, optionally supported on a catalyst support. In a preferred embodiment of the process of this invention, the paraffinic hydrocarbon is a paraffin selected from ethane, propane, mixtures of ethane and propane, naphtha, natural gas condensates, and mixtures of the aforementioned hydrocarbons; and the preferred olefins produced are ethylene, propylene, butylene, isobutylene, and butadiene.
In a preferred aspect of this invention, the feedstream comprising the paraffinic hydrocarbon and oxygen is preheated before introducing the feedstream into the autothermal oxidation reactor. The preheat temperature is greater than about 200xc2x0 C., but less than the temperature wherein reaction of the feedstream components begins. Preferably, the upper limit on the preheat temperature is less than about 900xc2x0 C.
In another preferred embodiment of this invention, the reactor comprises an exterior housing which holds the catalyst, the catalyst being provided in the form of a ceramic monolith support onto which the catalytic components, including the Group 8B metal and any promoter(s), have been deposited.
In another preferred aspect of this invention, the reactor comprises a modified fluidized bed characterized by an aspect ratio of less than about 1:1 in operating mode. As noted hereinbefore, the aspect ratio is the ratio of the height (depth) of the reactor to its cross-sectional dimension (diameter or width). In this reactor, the catalyst is provided typically in the form of spheres or granules.
In yet another preferred embodiment, the catalyst which is employed in the process of this invention comprises a Group 8B metal and at least one promoter supported on a catalytic support which has been pretreated with at least one support modifier. Preferably, the Group 8B metal is a platinum group metal. The preferred platinum group metal is platinum. The preferred promoter is selected from the elements of Groups 1B, 6B, 3A, 4A, 5A, (equivalent to Groups 11, 6, 13, 14, and 15), and mixtures of the aforementioned elements of the Periodic Table, as referenced by S. R. Radel and M. H. Navidi, in Chemistry, West Publishing Company, New York, 1990. The preferred support modifier is selected from Groups 1A, 2A, 3B, 4B, 5B, 6B, 1B, 3A, 4A, 5A (equivalent Groups 1, 2, 3, 4, 5, 6, 11, 13, 14, 15), and the lanthanide rare earths and actinide metals of the Periodic Table, as referenced by S. R. Radel and M. H. Navidi, ibid.
In a most preferred embodiment of the catalyst composition, the platinum group metal is platinum; the promoter is selected from tin, copper, and mixtures thereof; the support is selected from alumina, magnesia, and mixtures thereof; and the modifier is selected from tin, lanthanum, and mixtures thereof.
Any paraffinic hydrocarbon or mixture of paraffinic hydrocarbons can be employed in the process of this invention provided that an olefin, preferably, a mono-olefin, is produced. The term xe2x80x9cparaffinic hydrocarbon,xe2x80x9d as used herein, refers to a saturated hydrocarbon. Generally, the paraffin contains at least 2 carbon atoms. Preferably, the paraffin contains from 2 to about 25 carbon atoms, more preferably, from 2 to about 15 carbon atoms, and even more preferably, from 2 to about 10 carbon atoms. The paraffin can have a linear, branched, or cyclic structure, and can be a liquid or gas at ambient temperature and pressure. The paraffin can be supplied as an essentially pure paraffinic compound, or mixture of paraffinic compounds, or as a paraffin-containing mixture of hydrocarbons. Paraffin feeds which are suitably employed in the process of this invention include, but are not limited to, ethane, propane, butane, pentane, hexane, heptane octane, and higher homologues thereof, as well as complex higher boiling mixtures of paraffin-containing hydrocarbons, such as naphtha, gas oil, vacuum gas oil, and natural gas condensates. Additional feed components may include methane, nitrogen, carbon monoxide, carbon dioxide, and steam, if so desired. Minor amounts of unsaturated hydrocarbons may also be present. Most preferably, the paraffin is selected from ethane, propane, mixtures of ethane and propane, naphtha, natural gas condensates, and mixtures thereof.
In the process of this invention, the paraffinic hydrocarbon is contacted with an oxygen-containing gas. Preferably, the gas is molecular oxygen or molecular oxygen diluted with an unreactive gas, such as nitrogen, helium, carbon dioxide, or argon, or diluted with a substantially unreactive gas, such as carbon monoxide or steam. Any molar ratio of paraffin to oxygen is suitable provided the desired olefin is produced in the process of this invention. Preferably, the process is conducted fuel-rich and above the upper flammability limit. Generally, the molar ratio of paraffinic hydrocarbon to oxygen varies depending upon the specific paraffin feed and autothermal process conditions employed. Typically, the molar ratio of paraffinic hydrocarbon to oxygen ranges from about 3 to about 77 times the stoichiometric ratio of hydrocarbon to oxygen for complete combustion to carbon dioxide and water. Preferably, the molar ratio of paraffinic hydrocarbon to oxygen ranges from about 3 to about 13, more preferably, from about 4 to about 11, and most preferably, from about 5 to about 9 times the stoichiometric ratio of hydrocarbon to oxygen for complete combustion to carbon dioxide and water. These general limits are usually achieved by employing a molar ratio of paraffinic hydrocarbon to oxygen greater than about 0.1:1, preferably, greater than about 0.2:1, and by using a molar ratio of paraffinic hydrocarbon to oxygen usually less than about 3.0:1, preferably, less than about 2.7:1. For preferred paraffins, the following ratios are more specific. For ethane, the ethane to oxygen molar ratio is typically greater than about 1.5:1, and preferably, greater than about 1.8:1. The ethane to oxygen molar ratio is typically less than about 3.0:1, preferably, less than about 2.7:1. For propane, the propane to oxygen molar ratio is typically greater than about 0.9:1, preferably, greater than about 1.1:1. The propane to oxygen molar ratio is typically less than about 2.2:1, preferably, less than about 2.0:1. For naphtha, the naphtha to oxygen molar ratio is typically greater than about 0.3:1, preferably, greater than about 0.5:1. The naphtha to oxygen molar ratio is typically less than about 1.0:1, preferably, less than about 0.9:1.
When a high preheat temperature is used, for example, above 200xc2x0 C., the limits on the molar ratio of paraffinic hydrocarbon to oxygen can be shifted towards higher values. For example, at high preheat the molar ratio of paraffinic hydrocarbon to oxygen is typically greater than about 0.1:1 and less than about 4.0:1. Specifically, at high preheat the ethane to oxygen molar ratio is typically greater than about 1.5:1, preferably, greater than about 1.8:1, and typically less than about 4.0:1, preferably, less than about 3.2:1. At high preheat, the molar ratio of propane to oxygen is typically greater than about 0.9:1, preferably, greater than about 1.1:1, and typically, less than about 3.0:1, and preferably, less than about 2.6:1. At high preheat, the molar ratio of naphtha to oxygen is typically greater than about 0.3:1, preferably, greater than about 0.5:1, and typically, less than about 1.4:1, and preferably, less than about 1.3:1. As an advantageous feature of the process of this invention, hydrogen is co-fed with the paraffin and oxygen to the catalyst. The presence of hydrogen in the feedstream beneficially improves the conversion of hydrocarbon and the selectivity to olefins, while reducing the formation of deep oxidation products, such as, carbon monoxide and carbon dioxide. The molar ratio of hydrogen to oxygen can vary over any operable range provided that the desired olefin product is produced. Typically, the molar ratio of hydrogen to oxygen is greater than about 0.5:1, preferably, greater than about 0.7:1, and more preferably, greater than about 1.5:1. Typically, the molar ratio of hydrogen to oxygen is less than about 3.2:1, preferably, less than about 3.0:1, and more preferably, less than about 2.7:1.
At high preheat the molar ratio of hydrogen to oxygen typically is greater than about 0.1:1, preferably, greater than about 0.7:1, and more preferably, greater than about 1.5:1. At high preheat the molar ratio of hydrogen to oxygen is typically less than about 4.0:1, preferably, less than about 3.2:1, and more preferably, less than about 3.0:1.
Optionally, the feed may contain a diluent, which can be any gas or vaporizable liquid which does not interfere with the process of the invention. The diluent functions as a carrier of the reactants and products and facilitates the transfer of heat generated by the process. The diluent also helps to minimize undesirable secondary reactions and helps to expand the non-flammable regime for mixtures of the paraffin, hydrogen, and oxygen. Suitable diluents include nitrogen, argon, helium, carbon dioxide, carbon monoxide, methane, and steam. The concentration of diluent in the feed can vary over a wide range. If a diluent is used, the concentration of diluent is typically greater than about 0.1 mole percent of the total reactant feed including paraffin, oxygen, hydrogen, and diluent. Preferably, the amount of diluent is greater than about 1 mole percent of the total reactant feed. Typically, the amount of diluent is less than about 70 mole percent, and preferably, less than about 40 mole percent, of the total reactant feed.
The catalyst which is employed in the process of this invention beneficially comprises a Group 8B metal and at least one promoter, described hereinbelow, optionally supported on a catalyst support. The Group 8B metals include iron, cobalt, nickel, and the platinum group metals, namely, ruthenium, rhodium, palladium, osmium, iridium, and platinum. Mixtures of the aforementioned Group 8B metals may also be used. Preferably, the Group 8B metal is a platinum group metal; preferably, the platinum group metal is platinum. The catalyst also comprises at least one promoter, which is suitably defined as any element or elemental ion which is capable of enhancing the performance of the catalyst, as measured, for example, by an increase in the paraffin conversion, an increase in the selectivity to olefin, a decrease in the selectivities to deep oxidation products, such as carbon monoxide and carbon dioxide, and/or an increase in catalyst stability and lifetime. For the purposes of this invention, the term xe2x80x9cpromoterxe2x80x9d does not include t group metals. Preferably, the promoter is selected from the elements of Groups 1B (Cu, Ag, Au), 6B (Cr, Mo, W), 3A (for example, Al, Ga, In, Tl), 4A (for example, Ge, Sn, Pb), and 5A (for example, As, Sb, Bi), and mixtures thereof. More preferably, the promoter is selected from copper, tin, antimony, silver, indium, and mixtures thereof. Most preferably, the promoter is selected from copper, tin, antimony, and mixtures thereof.
Any atomic ratio of Group 8B metal to promoter can be employed in the catalyst, provided the catalyst is operable in the process of this invention. The optimal atomic ratio will vary with the specific Group 8B metal and promoter(s) employed. Generally, the atomic ratio of the Group 8B metal to promoter is greater than 0.10 (1:10), preferably, greater than about 0.13 (1:8), and more preferably, greater than about 0.17 (1:6). Generally, the atomic ratio of the Group 8B metal to promoter is less than about 2.0 (1:0.5), preferably, less than about 0.33 (1:3), and more preferably, less than about 0.25 (1:4). Although the promoter is used in a gram-atom amount equivalent to or greater than the Group 8B metal, the promoter nonetheless functions to enhance the catalytic effect of the catalyst. Compositions prepared with promoter alone, in the absence of Group 8B metal, are typically (but not necessarily always) catalytically inactive in the process. In contrast, the Group 8B metal is catalytically active in the absence of promoter, albeit with lesser activity.
The catalyst can be suitably employed in the form of a metallic gauze. More specifically, the gauze can comprise an essentially pure Group 8B metal or an alloy of Group 8B metals onto which the promoter is deposited. Suitable gauzes of this type include pure platinum gauze and platinum-rhodium alloy gauze coated with the promoter. The method used to deposit or coat the promoter onto the gauze can be any of the methods described hereinafter. Alternatively, a gauze comprising an alloy of a Group 8B metal and the promoter can be employed. Suitable examples of this type include gauzes prepared from platinum-tin, platinum-copper, and platinum-tin-copper alloys.
In another embodiment, the Group 8B metal and promoter are supported on a catalytic support. The loading of the Group 8B metal on the support can be any which provides for an operable catalyst in the process of this invention. In general, the loading of the Group 8B metal is greater than about 0.001 weight percent, preferably, greater than about 0.1 weight percent, and more preferably, greater than about 0.2 weight percent, based on the total weight of the Group 8B metal and support. Preferably, the loading of the Group 8B metal is less than about 80 weight percent, preferably, less than about 60 weight percent, and more preferably, less than about 10 weight percent, based on the total weight of the Group 8B metal and the support. Once the Group 8B metal loading is established, the desired atomic ratio of Group 8B metal to promoter determines the loading of the promoter.
The catalytic support comprises any material which provides a surface to carry the Group 8B metal, the promoter(s), and any support modifiers. Preferably, the support is thermally and mechanically stable under autothermal process conditions. Preferably, the catalytic support is a ceramic, such as, a refractory oxide, carbide, or nitride. Non-limiting examples of suitable ceramics include alumina, silica, silica-aluminas, aluminosilicates, including cordierite, magnesia, magnesium aluminate spinels, magnesium silicates, zirconia, titania, boria, zirconia toughened alumina (ZTA), lithium aluminum silicates, silicon carbide, oxide-bonded silicon carbide, and silicon nitride. Mixtures of the aforementioned refractory oxides, nitrides, and carbides may also be employed, as well as washcoats of the aforementioned materials on a support. Preferred ceramics include magnesia, alumina, silica, and amorphous or crystalline combinations of magnesia, alumina and silica, including mullite. Alpha (xcex1) and gamma (xcex3) alumina are preferred forms of alumina. Preferred combinations of alumina and silica comprise from about 65 to about 100 weight percent alumina and from essentially 0 to about 35 weight percent silica. Other refractory oxides, such as boria, can be present in smaller amounts in the preferred alumina and silica mixtures. Preferred zirconias include zirconia fully stabilized with calcia (FSZ) and zirconia partially stabilized with magnesia (PSZ), available from Vesuvius Hi-Tech Ceramics, Inc. Magnesia is the most preferred support, because it produces fewer cracking products and less carbon monoxide. Moreover, the hydrocarbon conversion and olefin selectivity tend to be higher with magnesia. The catalytic support may take a variety of shapes including that of porous or non-porous spheres, granules, pellets, irregularly shaped solid or porous particles, or any other shape which is suitable for catalytic reactors, including fixed bed, transport bed, and fluidized bed reactors. In a preferred form, the catalyst is a monolith. As used herein, the term xe2x80x9cmonolithxe2x80x9d means any continuous structure, including for example, honeycomb structures, foams, and fibers, including fibers woven into fabrics or made into non-woven mats or thin paper-like sheets. Monoliths do not, in general, contain significant microporosity. Foams have a sponge-like structure. More preferably, the support is a foam or fiber monolith. Fibers tend to possess higher fracture resistance as compared with foams and honeycombs. Preferred ceramic foams, available from Vesuvius Hi-Tech Ceramics, Inc., comprise magnesia, alpha alumina, zirconia, or mullite with a porosity ranging from about 5 to about 100 pores per linear inch (ppi) (2 to 40 pores per linear cm (ppcm)). Foams having about 45 ppi (18 ppcm) are more preferred. The term xe2x80x9cporosity,xe2x80x9d as used herein, refers to channel size or dimension. It is important to note that the foam supports are not substantially microporous structures. Rather, the foams are macroporous, meaning that they are low surface area supports with channels ranging in diameter from about 0.1 mm to about 5 mm. The foams are estimated to have a surface area less than about 10 m2/g, and preferably, less than about 2 m2/g, but greater than about 0.001 m2/g. Preferred ceramic fibers available from 3M Corporation as Nextel(trademark) brand ceramic fibers, typically have a diameter greater than about 1 micron (xcexcm), preferably, greater than about 5 xcexcm. The diameter is suitably less than about 20 xcexcm, preferably, less than about 15 xcexcm. The length of the fibers is generally greater than about 0.5 inch (1.25 cm), preferably, greater than about 1 inch (2.5 cm), and typically less than about 10 inches (25.0 cm), preferably, less than about 5 inches (12.5 cm). The surface area of the fibers is very low, being generally less than about 1 m2/g, preferably, less than about 0.3 m2/g, but greater than about 0.001 m2/g. Preferably, the fibers are not woven like cloth, but instead are randomly intertwined as in a mat or matted rug. Most preferred are Nextel(trademark) brand 440 fibers which consist of gamma alumina (70 weight percent), silica (28 weight percent), and boria (2 weight percent) and Nextel(trademark) brand 610 fibers which consist of alpha alumina (99 weight percent), silica (0.2-0.3 weight percent) and iron oxide (0.4-0.7 weight percent).
The deposition of the Group 8B metal and promoter(s) onto the support can be made by any technique known to those skilled in the art, for example, impregnation, ion-exchange, deposition-precipitation, vapor deposition, sputtering, and ion implantation. In one preferred methods the Group 8B metal is deposited onto the support by impregnation. Impregnation is described by Charles N. Satterfield in Heterogeneous Catalysis in Practice, McGraw-Hill Book Company, New York, 1980, 82-84, incorporated herein by reference. In this procedure, the support is wetted with a solution containing a soluble Group 8B compound, preferably, to the point of incipient wetness. The contacting temperature typically ranges from about ambient, taken as 23xc2x0 C., to about 100xc2x0 C., preferably, from about 23xc2x0 C. to about 50xc2x0 C. The contacting is conducted usually at ambient pressure. Non-limiting examples of suitable Group 8B compounds include the Group 8B nitrates, halides, sulfates, alkoxides, carboxylates, and Group 8B organometallic compounds, such as halo, amino, acetylacetonate, and carbonyl complexes. Preferably, the Group 8B compound is a platinum group halide, more preferably, a chloride, such as chloroplatinic acid. The solvent can be any liquid which solubilizes the Group 8B compound. Suitable solvents include water, aliphatic alcohols, aliphatic and aromatic hydrocarbons, and halo-substituted aliphatic and aromatic hydrocarbons. The concentration of the Group 8B compound in the solution generally ranges from about 0.001 molar (M) to about 10 M. After contacting the support with the solution containing the Group 8B compound, the support may be dried under air at a temperature ranging from about 23xc2x0 C. to a temperature below the decomposition temperature of the Group 8B compound, typically, a temperature between about 23xc2x0 C. and about 100xc2x0 C.
The deposition of the promoter can be accomplished in a manner analogous to the deposition of the Group 8B metal. Accordingly, if impregnation is used, then the support is wetted with a solution containing a soluble compound of the promoter at a temperature between about 23xc2x0 C. and about 100xc2x0 C., preferably, between about 23xc2x0 C. and about 50xc2x0 C., at about ambient pressure. Suitable examples of soluble promoter compounds include promoter halides, nitrates, alkoxides, carboxylates, sulfates, and organometallic compounds, such as amino, halo, and carbonyl complexes. Suitable solvents comprise water, aliphatic alcohols, aliphatic and aromatic hydrocarbons, and chloro-substituted aliphatic and aromatic hydrocarbons. Certain promoter compounds, such as compounds of antimony and tin, may be more readily solubilized in the presence of acid. For example, hydrochloric acid (5-25 weight percent) can be suitably employed. The concentration of the promoter compound in the solution generally ranges from about 0.01 M to about 10 M. Following deposition of the soluble promoter compound or mixture thereof, the impregnated support may be dried under air at a temperature between about 23xc2x0 C. and a temperature below the temperature wherein vaporization or decomposition of the promoter compound occurs. Typically, the drying is conducted at a temperature between about 23xc2x0 C. and about 100xc2x0 C.
In one method of preparing the catalyst, the Group 8B metal is deposited onto the support first, and thereafter the promoter is deposited onto the support. In an alternative method, the promoter is deposited first, followed by the deposition of the Group 8B metal. In a preferred method of preparing the catalyst, the Group 8B metal and the promoter are deposited simultaneously onto the support from the same deposition solution. In any of these methods, following one or more of the depositions, a calcination under oxygen is optional. If performed, the calcination is conducted at a temperature ranging from about 100xc2x0 C. to below the temperature at which volatilization of the metals becomes significant, typically, less than about 1,100xc2x0 C. Preferably, the calcination is conducted at a temperature between about 100xc2x0 C. and about 500xc2x0 C.
As a final step in the preparation of the catalyst, the fully-loaded support is reduced under a reducing agent, such as hydrogen, carbon monoxide, or ammonia, at a temperature between about 100xc2x0 C. and about 900xc2x0 C., preferably between about 125xc2x0 C. and about 800xc2x0 C., so as to convert the Group 8B metal substantially to its elemental form. The promoter may be reduced fully or partially, or not reduced at all, depending upon the specific promoter chosen and the reduction conditions. In addition, reduction at elevated temperatures may produce alloys of the Group 8B metal and the promoter. Alloys may provide enhanced catalyst stability by retarding vaporization of the promoter during the process of this invention.
In another preferred embodiment, the support is pretreated with a support modifier prior to loading the Group 8B and promoter(s). The support modifier can be any metal ion having a charge of +1 or greater. Preferably, the support modifier is selected from Groups 1A (Li, Na, K, Rb, Cs), 2A (for example, Mg, Ca, Sr, Ba), 3B (Sc, Y, La), 4B (Ti, Zr, Hf), 5B (V, Nb, Ta), 6B (Cr, Mo, W), 1B (Cu, Ag, Au), 3A (for example, Al, Ga, In), 4A (for example, Ge, Sn, Pb), 5A (for example, As, Sb, Bi), and the lanthanide rare earths (for example, Ce, Er, Lu, Ho) and actinide elements (specifically Th) of the Periodic Table previously identified. More preferably, the support modifier is selected from calcium, zirconium, tin, lanthanum, potassium, lutetium, erbium, barium, holmium, cerium, antimony, and mixtures thereof. Most preferably, the support modifier is selected from lanthanum, tin, antimony, calcium, and mixtures thereof. Certain elements, such as tin, antimony, and silver, may function as both promoter and support modifier simultaneously.
The procedure to modify the support comprises contacting the support with a solution containing a soluble compound of the support modifier. The contacting can involve ion-exchange or impregnation methods. Preferably, the modification procedure involves submerging the support in the solution such that essentially all of the surface area of the support is contacted with an excess of the solution. Compounds suitable for preparing the solution of support modifier include modifier nitrates, halides, particularly the chlorides, alkoxides, carboxylates, and organometallic complexes including amino, halo, alkyl, and carbonyl complexes. Suitable solvents include water, aliphatic alcohols, aromatic hydrocarbons, and halo-substituted aliphatic and aromatic hydrocarbons. Typically, the concentration of modifier compound in the solution ranges from about 0.001 M to about 10 M. Acidified solutions, for example, of hydrochloric acid and diluted solutions thereof, may be beneficially employed. The contact time generally ranges from about 1 minute to about 1 day. The contacting temperature suitably ranges from about 23xc2x0 C. to about 100xc2x0 C., and pressure is generally ambient. Alternatively, slurries of mixed oxides containing promoter and/or modifier elements, such as magnesium stannate (Mg2SnO4), can be deposited onto the support; The modified support is typically calcined, as noted hereinabove, or reduced under a reducing agent, such as hydrogen, at a temperature between about 100xc2x0 C. and about 900xc2x0 C., preferably, between about 200xc2x0 C. and about 800xc2x0 C. The choice of calcination or reduction depends on the element used to pretreat the support. If the element or its oxide is readily vaporizable, the pretreated support is reduced. If the element or its oxide is not readily vaporizable, then the pretreated support is calcined. As a guideline, the words xe2x80x9creadily vaporizablexe2x80x9d may be taken to mean that greater than about 1 weight percent of any metal component in the catalyst is vaporized in a period of about 24 hours under calcination conditions at about 200xc2x0 C. The term xe2x80x9creadily vaporizablexe2x80x9d may be given a narrower or broader definition, as desired.
Following the pretreatment modification, the Group 8B metal and promoter(s) are loaded onto the support. Then, the support is reduced as described hereinbefore Alternatively, the metal-loaded support may be calcined first and then reduced. Whether the modified support is calcined or not depends again upon the vaporization potential of the modifier metal(s) and promoter(s) employed. Supports modified with metals or metal oxides which tend to vaporize readily are typically not calcined. Support modified with metals or metal oxides which do not vaporize readily can be calcined.
The process of this invention is advantageously conducted under autothermal process conditions. The term xe2x80x9cautothermal process conditionsxe2x80x9d means that the heat generated by reaction of the feed is sufficient to support the catalytic process which converts the paraffin to the olefin. Accordingly, the need for an external heating source to supply the energy for the process can be eliminated. In order to maintain autothermal conditions, the catalysts of the prior art are required to support combustion beyond the normal, fuel-rich limit of flammability. This is not a requirement in the present invention. Here, autothermal conditions can also be maintained with a catalyst which does not support combustion beyond the normal, fuel-rich limit of flammability, provided that hydrogen and optionally a preheat are supplied to the process.
Ignition can be effected by preheating the feed to a temperature sufficient to effect ignition when contacted with the catalyst. Alternatively, the feed can be ignited with an ignition source, such as a spark or flame. Upon ignition, the reaction-generated heat causes the temperature to take a step change jump to a new steady state level that is herein referred to as the autothermal reaction.
While running autothermal, the paraffin feed does not have to be preheated, so long as the feed contains hydrogen or the catalyst supports combustion beyond the normal, fuel-rich limit of flammability. (The word xe2x80x9ccombustion,xe2x80x9d as used herein, means the reaction of the hydrocarbon with oxygen unaided by hydrogen.) Preheating the feedstream, however, has certain advantages. The advantages comprise a decrease in oxygen and hydrogen consumed, an increase in the paraffin concentration in the feed, an increase in the operating paraffin to oxygen molar ratio, and a net increase in recycle hydrogen in the product cream. In addition, catalysts can be used which do not support combustion beyond the normal fuel-rich limit of flammability. These advantages are particularly significant when the preheating is conducted at a temperature greater than about 200xc2x0 C. and less than the temperature wherein reaction of the feedstream components begins. Suitable preheat temperatures are typically greater than about 40xc2x0 C., preferably, greater than about 125xc2x0 C., and even more preferably, greater than about 200xc2x0 C. In another preferred embodiment, the preheat temperature is greater than about 400xc2x0 C. Suitable preheat temperatures are typically less than about 900xc2x0 C., preferably, less than about 800xc2x0 C., and more preferably less than about 600xc2x0 C.
As a general rule, the autothermal process operates at close to the adiabatic temperature (that is, essentially without loss of heat), which is typically greater than about 75xc2x0 C., and preferably, greater than about 925xc2x0 C. Typically, the autothermal process operates at a temperature less than about 1,150xc2x0 C., and preferably, less than about 1,050xc2x0 C. Optionally, the temperature at the reactor exit can be measured, for example, by using a Pt/Ptxe2x80x94Rh thin wire thermocouple. With a monolith catalyst, the thermocouple can be sandwiched between the monolith and the downstream radiation shield. Measurement of temperature close to the reactor exit may be complicated by the high temperature involved and the fragility of the thermocouple. Thus, as an alternative, one skilled in the art can calculate the adiabatic temperature at the reactor exit from a knowledge of the preheat temperature and the exit stream composition. The xe2x80x9cadiabatic temperaturexe2x80x9d is the temperature of the product stream without any heat loss, that is, when all of the heat generated by the process is used to heat the products. Typically, the measured temperature is found to be within about 25xc2x0 C. of the calculated adiabatic temperature.
The operating pressure is typically equal to or greater than about 1 atmosphere absolute (atm abs) (100 kPa abs). Typically, the pressure is less than about 20 atm abs (2,000 kPa abs), preferably, less than about 10 atm (1,000 kPa abs), and more preferably, less than about 7 atm abs (700 kPa abs).
Since the products of this process must be removed rapidly from the reaction zone, gas hourly space velocities are very high. The specific gas hourly space velocity employed will depend upon the choice of reactor cross sectional dimension (for example, diameter) and the form and weight of the catalyst particles. Generally, the gas hourly space velocity (GHSV), calculated as the total flow of the hydrocarbon, oxygen, hydrogen, and optional diluent flows, is greater than about 50,000 ml total feed per ml catalyst per hour (hxe2x88x921) measured at standard temperature and pressure (0xc2x0 C., 1 atm) (STP). Preferably, the GHSV is greater than about 80,000 hxe2x88x921, and more preferably, greater than 100,000 hxe2x88x921. Generally, the gas hourly space velocity is less than about 6,000,000 hxe2x88x921, preferably, less than about 4,000,000 hxe2x88x921, more preferably, less than 3,000,000 hxe2x88x921, measured as the total flow at STP. Gas flows are typically monitored in units of liters per minute at standard temperature and pressure (slpm). The conversion gas flow from xe2x80x9cslpmxe2x80x9d units to gas hourly s"" pace velocity units (hxe2x88x921) is made as follows:       GHSV    ⁢          xe2x80x83        ⁢          h              -        1              =            slpm      xc3x97      1000      ⁢              xe2x80x83            ⁢              cm        3            ⁢              /            ⁢      min      xc3x97      60      ⁢              xe2x80x83            ⁢      min      ⁢              /            ⁢      h              cross      ⁢              -            ⁢      sectional      ⁢              xe2x80x83            ⁢      area      ⁢              xe2x80x83            ⁢      of      ⁢              xe2x80x83            ⁢      catalyst      ⁢              xe2x80x83            ⁢              (                  cm          2                )            xc3x97      length      ⁢              xe2x80x83            ⁢              (        cm        )            
The residence time of the reactants in the reactor is simply calculated as the inverse of the gas hourly space velocity. At the high space velocities employed in the process of this invention, the residence time is on the order of milliseconds. Thus, for example, a gas hourly space velocity of 100,000 hxe2x88x921 measured at STP is equivalent to a residence time of 36 milliseconds at STP.
The process of this invention may be conducted in any reactor designed for use under adiabatic, autothermal process conditions. In one preferred design, the catalyst is prepared on a monolith support which is sandwiched between two radiation shields inside a reactor housing. Alternatively, fixed bed and fluidized bed reactors can be used with catalysts in the form of pellets, spheres, and other particulate shapes. Continuous and intermittent flow of the feedstream are both suitable. It is noted that fluidized bed reactors of the prior art typically possess an aspect ratio in static mode of greater than 1:1, and more preferably, greater than about 5:1. Static mode is defined as the unfluidized or fixed bed configuration. Fluidized bed reactors are generally operated in a bubbling, turbulent, or fast-fluidized regime with expanded beds measuring from about 1.5 to 15 times the static depth. Typically, the aspect ratio in operating mode is greater than about 5:1 to 10:1. For full fluidization, a catalyst particle size ranging between about 30 and 1,000 microns is satisfactory.
It is believed that the oxidation reaction of this invention occurs predominantly at the reactor entry, which in the case of a stationary catalyst is at the front edge of the catalyst. Such a theory should not be binding or limiting of the invention in any manner. In view of this theory, the optimal reactor for the process of this invention should possess a large cross-sectional dimension and a short height (or depth). On a commercial scale, for example, a catalyst bed of diameter about 5 to 8 feet (1.5 m to 2.4 m) and a height of about 1 inch (2.5 cm) may be suitably employed. Additionally, it is believed that catalyst located at the front edge of a stationary bed can deactivate more quickly with time. As a consequence, longer catalyst lifetime and better selectivities can be achieved by circulating particles of the catalyst, rather than using a stationary bed.
A preferred reactor design for the process of this invention comprises a modified fluidized bed reactor, characterized in that its aspect ratio in operating mode, and preferably also in static mode (unfluidized or fixed bed configuration), is less than 1:1, and more preferably, less than about 0.1:1, but greater than about 0.001:1. Most preferably, the aspect ratio is about 0.01:1. This unique fluidized bed is operated above the minimum fluidization flow with an expanded bed on the order of about 2 or 3 times the static depth, and preferably, less than about 1.5 times the static depth. For the purposes of this invention, xe2x80x9cminimum fluidization flowxe2x80x9d is defined as the minimum gas velocity at which the catalyst particles are suspended under operating conditions. The velocity necessary to achieve minimum fluidization depends upon the density and viscosity of the gas phase and the catalyst particle size and density. One skilled in the art would know how to calculate the minimum fluidization flow for any given gas composition and catalyst particle. A suitable authority on the subject is found in Fluidization Engineering, by D. Kunii and O. Levenspeil, 2nd ed., Butterworth-Heineman, 1989, incorporated herein by reference. A catalyst particle size of between about 500 and about 850 microns (23-30 US mesh) is suitable for feed velocities of about 0.05 to 5 meters per second (mps) at standard temperature and pressure. An advantage of the modified fluidized bed reactor may result from its continuous circulation (fluidization), which results in continuous renewal of catalyst particles at the reactor entry. This configuration produces substantially better product yields than a stationary catalyst.
When a paraffinic hydrocarbon is contacted with oxygen under autothermal process conditions in the presence of a co-feed of hydrogen and in the presence of the multi-metallic catalyst described hereinabove, an olefin, preferably a mono-olefin, is produced. Ethane is converted primarily to ethylene. Propane and butane are converted primarily to ethylene and propylene. Isobutane is converted primarily to isobutylene and propylene. Naphtha and other higher molecular weight paraffins are converted primarily to ethylene and propylene.
The conversion of paraffinic hydrocarbon in the process of this invention can vary depending upon the specific feed composition, catalyst composition, reactor, and process conditions employed. For the purposes of this invention, xe2x80x9cconversionxe2x80x9d is defined as the mole percentage of paraffinic hydrocarbon in the feed which is converted to products. Generally, at constant pressure and space velocity, the conversion increases with increasing temperature. Typically, at constant temperature and pressure, the conversion does not change significantly over a wide range of high space velocities employed. In this process, the conversion of paraffinic hydrocarbon is typically greater than about 50 mole percent, preferably, greater than about 60 mole percent, and more preferably, greater than about 70 mole percent.
Likewise, the selectivity to products will vary depending upon the specific feed composition, catalyst composition, reactor, and process conditions employed. For the purposes of this invention, xe2x80x9cselectivityxe2x80x9d is defined as the percentage of carbon atoms in the converted paraffin feed which react to form a specific product. For example, the olefin selectivity is calculated as follows:                               Moles          ⁢                      xe2x80x83                    ⁢          of          ⁢                      xe2x80x83                    ⁢          olefin          ⁢                      xe2x80x83                    ⁢          formed          xc3x97                                              Number          ⁢                      xe2x80x83                    ⁢          of          ⁢                      xe2x80x83                    ⁢          carbon          ⁢                      xe2x80x83                    ⁢          atoms          ⁢                      xe2x80x83                    ⁢          in          ⁢                      xe2x80x83                    ⁢          olefin          xc3x97          100                                                  Moles          ⁢                      xe2x80x83                    ⁢          of          ⁢                      xe2x80x83                    ⁢          paraffin          ⁢                      xe2x80x83                    ⁢          converted          xc3x97                                              Number          ⁢                      xe2x80x83                    ⁢          of          ⁢                      xe2x80x83                    ⁢          carbon          ⁢                      xe2x80x83                    ⁢          atoms          ⁢                      xe2x80x83                    ⁢          in          ⁢                      xe2x80x83                    ⁢          paraffin                    
Generally, the olefin selectivity increases with increasing temperature up to a maximum value and declines as the temperature continues to rise. Usually, the olefin selectivity does not change substantially over a wide range of high space velocities employed. In the process of this invention, the olefin selectivity is typically greater than about 50 carbon atom percent, preferably, greater than about 60 carbon atom percent, more preferably, greater than about 70 carbon atom percent, and even more preferably, greater than about 80 carbon atom percent. Other products formed in smaller quantities include methane, carbon monoxide, carbon dioxide, propane, butenes, butadiene, propadiene, aceylene, methylacetylene, and C6+ hydrocarbons. Acetylene can be hydrogenated to ethylene downstream to increase the overall selectivity to olefin. At least part of the carbon monoxide, carbon dioxide, and methane formed may be recycled to the reactor.
Water is also formed in the process of this invention from the reaction of hydrogen or hydrocarbon. The presence of hydrogen in the feed minimizes the formation of carbon oxides by reacting with the oxygen to produce water and energy. Accordingly, it is advantageous to recycle the hydrogen in the product stream, obtained from the dehydrogenation of the paraffin, back to the reactor. Optimally, the hydrogen needed to meet the demands of the process essentially equals the hydrogen formed during conversion of the paraffin to olefin. Under these balanced conditions, the hydrogen forms a closed loop wherein there is essentially no demand for additional hydrogen to be added to the feed. Such conditions are more easily met when the feed is preheated and a higher hydrocarbon to oxygen molar ratio is employed.